Hydrocracking catalyst containing beta and Y zeolites, and process for its use to produce naphtha

ABSTRACT

Increased yields of naphtha and increased catalyst activity are obtained in a hydrocracking process by the use of a catalyst containing a beta zeolite and a Y zeolite having a unit cell size from 24.38 to 24.50 angstrom. The catalyst has a relatively high amount of Y zeolite relative to beta zeolite.

FIELD OF THE INVENTION

The invention relates to catalyst compositions and their use inhydrocarbon conversion processes, particularly hydrocracking. Theinvention more specifically relates to a catalyst composition thatcomprises a Y zeolite and a beta zeolite as active cracking components.The invention specifically relates to a hydrocracking process thatproduces naphtha.

BACKGROUND OF THE INVENTION

Petroleum refiners often produce desirable products such as naphtha andgasoline as well as higher boiling liquids, such as turbine fuel, dieselfuel, and other hydrocarbon liquids known as middle distillates, byhydrocracking a hydrocarbon feedstock derived from crude oil.Hydrocracking also has other beneficial results such as removing sulfurand nitrogen from the feedstock by hydrotreating. Feedstocks most oftensubject to hydrocracking are gas oils and heavy gas oils recovered fromcrude oil by distillation.

Hydrocracking is generally carried out by contacting, in an appropriatereactor vessel, the gas oil or other hydrocarbon feedstock with asuitable hydrocracking catalyst under appropriate conditions, includingan elevated temperature and an elevated pressure and the presence ofhydrogen so as to yield a lower overall average boiling point productcontaining a distribution of hydrocarbon products desired by therefiner. Although the operating conditions within a hydrocrackingreactor have some influence on the yield of the products, thehydrocracking catalyst is a prime factor in determining such yields.

Hydrocracking catalysts are subject to initial classification on thebasis of the nature of the predominant cracking component of thecatalyst. This classification divides hydrocracking catalysts into thosebased upon an amorphous cracking component such as silica-alumina andthose based upon a zeolitic cracking component such as beta or Yzeolite. Hydrocracking catalysts are also subject to classification onthe basis of their intended predominant product of which the two mainproducts are distillate and naphtha, a term which in the hydrocrackingrefining art refers to distillable petroleum derived fractions having aboiling point range that is below that of distillate. Naphtha typicallyboils in the range of from the boiling point of hydrocarbons having sixcarbon atoms per molecule (i.e., C₆) to 216° C. (420° F.) and includesthe product recovered at a refinery as gasoline. At the present time,naphtha and gasoline are in high demand. For this reason, refiners havebeen focusing on hydrocracking catalysts that selectively produce anaphtha fraction.

The three main catalytic properties by which the performance of ahydrocracking catalyst for producing naphtha is evaluated are activity,selectivity, and stability. Activity may be determined by comparing thetemperature at which various catalysts must be utilized under otherwiseconstant hydrocracking conditions with the same feedstock so as toproduce a given percentage, normally about 65 percent, of productsboiling below 216° C. (420° F.). The lower the temperature required fora given catalyst, the more active such a catalyst is in relation to acatalyst requiring a higher temperature. Selectivity of hydrocrackingcatalysts for producing naphtha may be determined during the foregoingdescribed activity test and is measured as a percentage of the fractionof the product boiling in the desired naphtha product range, e.g., C₆ to216° C. (420° F.). Stability is a measure of how well a catalystmaintains its activity over an extended time period when treating agiven hydrocarbon feedstock under the conditions of the activity test.Stability is generally measured in terms of the change in temperaturerequired per day to maintain a 65 percent or other given conversion.

Although cracking catalysts for producing naphtha are known and used incommercial environments, there is always a demand for new hydrocrackingcatalysts with superior overall activity, selectivity, and stability forproducing naphtha.

BRIEF SUMMARY OF THE INVENTION

It has been found that hydrocracking catalysts containing a beta zeolitehaving an overall silica to alumina (SiO₂ to Al₂O₃) mole ratio of lessthan 30 and a SF₆ adsorption capacity of at least 28 weight-percent(hereinafter wt-%) and also containing a Y zeolite having a unit cellsize or dimension a_(o), from 24.38 to 24.50 angstrom, wherein thecatalyst has a weight ratio of Y zeolite to beta zeolite of from 5 to 12on a dried basis, have substantially improved activity and selectivitycompared to other hydrocracking catalysts now commercially available foruse in hydrocracking processes for producing naphtha. The catalyst alsocontains a metal hydrogenation component such as nickel, cobalt,tungsten, molybdenum, or any combination thereof.

It is believed that a hydrocracking catalyst containing such a Y zeoliteand such a beta zeolite is novel to the art.

Under typical hydrocracking conditions, including elevated temperatureand pressure and the presence of hydrogen, such catalysts are highlyeffective for converting gas oil and other hydrocarbon feedstocks to aproduct of lower average boiling point and lower average molecularweight. The product contains a relatively large proportion of componentsboiling in the naphtha range, which as defined herein is from C₆ to 216°C. (420° F.).

INFORMATION DISCLOSURE

Beta and Y zeolites have been proposed in combination as components ofseveral different catalysts including catalysts for hydrocracking. Forinstance, U.S. Pat. No. 5,275,720 describes a hydrocracking processusing a catalyst comprising a beta zeolite and a dealuminated Y zeolitehaving an overall silica to alumina mole ratio greater than 6.0 and aunit cell size between about 24.40 and 24.65 angstrom. The weight ratioof the dealuminated Y zeolites to the beta zeolite may be in the rangeof 0.25 to 4. Preferred Y zeolites include LZ-210 zeolite.

U.S. Pat. No. 5,279,726 describes a hydrocracking process using acatalyst comprising a beta zeolite and a Y zeolite having a unit cellsize greater than 24.40 angstrom, and usually a water vapor sorptivecapacity at 25° C. (77° F.) and a p/p_(o) value of 0.10 of greater than15 weight percent. The overall silica to alumina mole ratio of themodified Y zeolites generally ranges between 5.1 and 6.0, although themodified Y zeolites may have a silica-to-alumina mole ratio above 6.0,e.g., between 6 and 20. Preferred Y zeolites include LZY-84 or Y-84zeolite. The weight ratio of the Y zeolites to the beta zeolite may bein the range of 0.33 to 3.

U.S. Pat. No. 5,350,501 describes a hydrocracking process using acatalyst comprising a support comprising a beta zeolite and a Y zeolitehaving either (1) a unit cell size less than about 24.45 angstrom or (2)a sorptive capacity for water at 25° C. (77° F.) and a p/p_(o) value of0.10 of less than 10.00 weight percent. LZ-10 is a preferred Y zeolite.The weight ratio of the Y zeolites to the beta zeolite may be in therange of 0.33 to 3.

DETAILED DESCRIPTION OF THE INVENTION

The process and composition disclosed herein may be used to convert afeedstock containing organic compounds into products, particularly byacid catalysis, such as hydrocracking organic compounds especiallyhydrocarbons into a product of lower average boiling point and loweraverage molecular weight. The composition, which may be a catalystand/or a catalyst support, comprises a beta zeolite and a Y zeolite. Thecomposition may also comprise a refractory inorganic oxide. When used asa catalyst for hydrocracking, the composition contains a beta zeolite, aY zeolite, a refractory inorganic oxide, and a hydrogenation component.

The hydrocracking process and composition disclosed herein centers onusing a catalyst containing a particular beta zeolite and a particular Yzeolite at a relatively high weight ratio of Y zeolite to beta zeolite.The beta zeolite in some embodiments has a relatively low silica toalumina mole ratio and a relatively high SF₆ adsorption capacity. The Yzeolite has a relatively low silica to alumina mole ratio and arelatively high unit cell size. It has been found that differingperformance results when such a beta zeolite and such a Y zeolite areincorporated in a hydrocracking catalysts in this way. Not only is theactivity of the hydrocracking catalyst higher than that of catalystscontaining the Y zeolite, but the yield of product boiling in thenaphtha range is higher too. The low amount of beta zeolite relative toY zeolite improves the yield of naphtha. Furthermore, the stability ofthe hydrocracking catalyst is higher than that of catalysts having alower weight ratio of Y zeolite to beta zeolite. The low amount of betazeolite relative to Y zeolite helps to decrease the formation ofundesirable heavy polynuclear aromatic byproducts that may decreasestability.

Beta zeolite is well known in the art as a component of hydrocrackingcatalysts. Beta zeolite is described in U.S. Pat. No. 3,308,069 and ReNo. 28,341, which are hereby incorporated by reference herein in theirentireties. The beta zeolite that is used in the process and compositiondisclosed herein has a silica to alumina mole ratio of less than 30 inone embodiment, less than 25 in another embodiment, more than 9 and lessthan 30 in yet another embodiment, more than 9 and less than 25 in afurther embodiment, more than 20 and less than 30 in another embodiment,or more than 15 and less than 25 in still another embodiment. As usedherein, unless otherwise indicated, the silica to alumina (SiO₂ toAl₂O₃) mole ratio of a zeolite is the mole ratio as determined on thebasis of the total or overall amount of aluminum and silicon (frameworkand non-framework) present in the zeolite, and is sometimes referred toherein as the overall silica to alumina (SiO₂ to Al₂O₃) mole ratio.

Beta zeolite is usually synthesized from a reaction mixture containing atemplating agent. The use of templating agents for synthesizing betazeolite is well known in the art. For example, U.S. Pat. No. 3,308,069and Re No. 28,341 describe using tetraethylammonium hydroxide and U.S.Pat. No. 5,139,759, which is hereby incorporated herein by reference inits entirety, describes using the tetraethylammonium ion derived fromthe corresponding tetraethylammonium halide. Another standard method ofpreparing beta zeolite is described in the book titled VerifiedSynthesis of Zeolitic Materials, by H. Robson (editor) and K. P.Lillerud (XRD Patterns), second revised edition, ISBN 0-444-50703-5,Elsevier, 2001. It is believed that the choice of a particulartemplating agent is not critical to the success of the process disclosedherein. In one embodiment the beta zeolite is calcined in air at atemperature of from 500 to 700° C. (932 to 1292° F.) for a timesufficient to remove to remove the templating agent from the betazeolite. Calcination to remove the templating agent can be done beforeor after the beta zeolite is combined with the support and/or thehydrogenation component. Although it is believed that the templatingagent could be removed at calcination temperatures above 700° C. (1292°F.), very high calcination temperatures could significantly decrease theSF₆ adsorption capacity of beta zeolite. For this reason it is believedthat calcination temperatures above 750° C. (1382° F.) for removing thetemplating agent should be avoided when preparing the beta zeolite foruse in the process disclosed herein. It is critical to the processdisclosed herein that the SF₆ adsorption capacity of the beta zeolite isat least 28 wt-%.

While it is known that steaming a zeolite such as beta results inchanges to the actual crystalline structure of the zeolite, theabilities of present day analytical technology have not made it possibleto accurately monitor and/or characterize these changes in terms ofimportant structural details of the zeolite. Instead, measurements ofvarious physical properties of the zeolite such as surface area are usedas indicators of changes that have occurred and the extent of thechanges. For instance, it is believed that a reduction in the zeolite'scapacity to adsorb sulfur hexafluoride (SF₆) after being steamed iscaused by a reduction in the crystallinity of the zeolite or in the sizeor accessibility of the zeolite's micropores. It is, however, anindirect correlation of the changes in the zeolite that may beundesirable, since the SF₆ adsorption capacity in the catalyst used inthe process and composition disclosed herein is relatively high. Inembodiments of the process and composition disclosed herein, the SF₆adsorption capacity of the beta zeolite, whether steam treated or not,should be at least 28 wt-%.

Accordingly, the beta zeolite of the process and composition disclosedherein may be characterized in terms of SF₆ adsorption. This is arecognized technique for the characterization of microporous materialssuch as zeolites. It is similar to other adsorption capacitymeasurements, such as water capacity, in that it uses weight differencesto measure the amount of SF₆ which is adsorbed by a sample which hasbeen pretreated to be substantially free of the adsorbate. SF₆ is usedin this test since because its size and shape hinders its entrance intopores having a diameter of less than about 6 angstrom. It thus can beused as one measurement of available pore mouth and pore diametershrinkage. This in turn is a measurement of the effect of steaming onthe zeolite. In a simplistic description of this measurement method, thesample is preferably first predried in a vacuum at 300° C. (572° F.) forone hour, then heated at atmospheric pressure in air at 650° C. (1202°F.) for two hours, and finally weighed. It is then exposed to the SF₆for one hour while the sample is maintained at a temperature of 20° C.(68° F.). The vapor pressure of the SF₆ is maintained at that providedby liquid SF₆ at 400 torr (53.3 kPa (7.7 psi)). The sample is againweighed to measure the amount of adsorbed SF₆. The sample may besuspended on a scale during these steps to facilitate these steps.

In any mass production procedure involving techniques such as steamingand heating there is a possibility for individual particles to besubjected to differing levels of treatment. For instance, particles onthe bottom of a pile moving along a belt may not be subjected to thesame atmosphere or temperature as the particles which cover the top ofthe pile. This factor must be considered during manufacturing and alsoduring analysis and testing of the finished product. It is, therefore,recommended that any test measure done on the catalyst is performed on arepresentative composite sample of the entire quantity of finishedproduct to avoid being misled by measurements performed on individualparticles or on a non-representative sample. For instance, an adsorptioncapacity measurement is made on a representative composite sample.

Although the process and the composition disclosed herein can use a betazeolite that has not been subjected to a steaming treatment, the processand the composition disclosed herein can also use beta zeolite that issubjected to steaming, provided that the steaming is relatively mild incomparison to steaming of beta zeolite in the literature. Under theproper conditions and for the proper time, steaming beta zeolite canyield a catalyst that can be used in the process and compositiondisclosed herein.

Hydrothermally treating zeolites for use in hydrocracking catalysts is arelatively blunt tool. For any given zeolite, steaming decreases theacidity of the zeolite. When the steamed zeolite is used as ahydrocracking catalyst, the apparent result is that the overall naphthayield increases but the catalyst's activity decreases. This apparenttradeoff between yield and activity has meant that to achieve highactivity means not to steam the beta zeolite, but at the expense oflower product yields. This apparent tradeoff between naphtha yield andactivity must be considered and is a limit to the improvement thatappears to be obtainable by steaming the beta zeolite. When the steamedbeta zeolite is used in the catalysts disclosed herein, the improvementin activity over catalysts containing only Y zeolite would appearlimited while the improvement in naphtha yield over such catalysts wouldappear more enhanced.

If the beta zeolite is to be steamed, such steaming can be performedsuccessfully in different ways, with the method which is actuallyemployed commercially often being greatly influenced and perhapsdictated by the type and capability of the available equipment. Steamingcan be performed with the beta zeolite retained as a fixed mass or withthe beta zeolite being confined in a vessel or being tumbled whileconfined in a rotating kiln. The important factors are uniform treatmentof all beta zeolite particles under appropriate conditions of time,temperature and steam concentration. For instance, the beta zeoliteshould not be placed such that there is a significant difference in theamount of steam contacting the surface and the interior of the betazeolite mass. The beta zeolite may be steam treated in an atmospherehaving live steam passing through the equipment providing low steamconcentration. This may be described as being at a steam concentrationof a positive amount less than 50 mol-%. Steam concentrations may rangefrom 1 to 20 mol-% or from 5 to 10 mol-%, with small-scale laboratoryoperations extending toward higher concentrations. The steaming may beperformed for a positive time period of less than or equal to 1 or 2hours or for 1 to 2 hours at a temperature of less than or equal toabout 600° C. (1112° F.) at atmospheric pressure and a positive contentof steam of less than or equal to 5 mol-%. The steaming may be performedfor a positive time period of less than or equal to 2 hours at atemperature of less than or equal to about 650° C. (1202° F.) atatmospheric pressure and a positive content of steam of less than orequal to 10 mol-%. The steam contents are based on the weight of vaporscontacting the beta zeolite. Steaming at temperatures above 650° C.(1202° F.) appears to result in beta zeolite that is not useful in theprocess disclosed herein since the SF₆ adsorption capacity of theresulting beta zeolite is too low. Temperatures below 650° C. (1202° F.)can be used, and the steaming temperature can be from about 600° C.(1112° F.) to about 650° C. (1202° F.), or less than 600° C. (1112° F.).It is taught in the art that there is normally an interplay between timeand temperature of steaming, with an increase in temperature reducingthe required time. Nevertheless, if steaming is done, for good resultsit appears a time period of about ½ to about 2 hours or about 1 to about1½ hours can be used. The method of performing steaming on a commercialscale may be by means of a rotary kiln having steam injected at a ratewhich maintains an atmosphere of about 10 mol-% steam.

An exemplary lab scale steaming procedure is performed with the zeoliteheld in a 6.4 cm (2½ inch) quartz tube in a clam shell furnace. Thetemperature of the furnace is slowly ramped up by a controller. Afterthe temperature of the zeolite reaches 150° C. (302° F.) steam generatedfrom deionized water held in a flask is allowed to enter the bottom ofthe quartz tube and pass upward. Other gas can be passed into the tubeto achieve the desired steam content. The flask is refilled as needed.In the exemplary procedure the time between cutting in the steam and thezeolite reaching 600° C. (1112° F.) is about one hour. At the end of theset steam period the temperature in the furnace is reduced by resettingthe controller to 20° C. (68° F.). The furnace is allowed to cool to400° C. (752° F.) (about 2 hours) and the flow of steam into the quartztube is stopped. The sample is removed at 100° C. (212° F.) and placedin a lab oven held overnight at 110° C. (230° F.) with an air purge.

The beta zeolite of the process and composition disclosed herein is nottreated with an acid solution to effect dealumination. In this regard itis noted that essentially all raw (as synthesized) beta zeolite isexposed to an acid to reduce the concentration of alkali metal (e.g.,sodium) which remains from synthesis. This step in the beta zeolitemanufacture procedure is not considered part of the treatment ofmanufactured beta zeolite as described herein. In one embodiment, duringthe treatment and catalyst manufacturing procedures, the beta zeolite isexposed to an acid only during incidental manufacturing activities suchas peptization during forming or during metals impregnation. In anotherembodiment, the beta zeolite is not acid washed after the steamingprocedure as to remove aluminum “debris” from the pores.

Also included in the process and composition disclosed herein is a Yzeolite having a unit cell size from 24.38 to 24.50 angstrom. In oneembodiment, the Y zeolite has an overall silica to alumina mole ratiofrom 5.0 to 11.0. The term “Y zeolite” as used herein is meant toencompass all crystalline zeolites having either the essential X-raypowder diffraction pattern set forth in U.S. Pat. No. 3,130,007 or amodified Y zeolite having an X-ray powder diffraction pattern similar tothat of U.S. Pat. No. 3,130,007 but with the d-spacings shifted somewhatdue, as those skilled in the art will realize, to cation exchanges,calcinations, etc., which are generally necessary to convert the Yzeolite into a catalytically active and stable form. The process andcomposition disclosed herein require a Y zeolite having either or bothof the two properties mentioned above, such Y zeolites being modified Yzeolites in comparison to the Y zeolite taught in U.S. Pat. No.3,130,007. As used herein, unit cell size means the unit cell size asdetermined by X-ray powder diffraction.

The Y zeolites used in the process and composition disclosed herein arelarge pore zeolites having an effective pore size greater than 7.0angstrom. Since some of the pores of the Y zeolites are relativelylarge, the Y zeolites allow molecules relatively free access to theirinternal structure. The pores of the Y zeolites permit the passagethereinto of benzene molecules and larger molecules and the passagetherefrom of reaction products.

One group of Y zeolites that may be used in the process and compositiondisclosed herein includes zeolites that are sometimes referred to asultrastable Y zeolites. The composition and properties of this group ofY zeolites are, in essence, prepared by a four step procedure. First, aY zeolite in the alkali metal form (usually sodium) and typically havinga unit cell size of about 24.65 angstrom is cation exchanged withammonium ions. The ammonium exchange step typically reduces the sodiumcontent of the starting sodium Y zeolite from a value usually greaterthan about 8 wt-%, usually from about 10 to about 13 wt-%, calculated asNa₂O, to a value in the range from about 0.6 to about 5 wt-%, calculatedas Na₂O. Methods of carrying out the ion exchange are well known in theart.

Second, the Y zeolite from the first step is calcined in the presence ofwater vapor. For example, the Y zeolite is calcined in the presence ofat least 1.4 kpa(absolute) (hereinafter kPa(a)) (0.2 psi(absolute)(hereinafter psi(a))), at least 6.9 kpa(a) (1.0 psi(a)), or at least 69kpa(a) (10 psi(a)) water vapor, in three embodiments. In two otherembodiments, the Y zeolite is calcined in an atmosphere consistingessentially of or consisting of steam. The Y zeolite is calcined so asto produce a unit cell size in the range of 24.40 to 24.64 angstrom.

Third, the Y zeolite from the second step is ammonium exchanged onceagain. The second ammonium exchange further reduces the sodium contentto less than about 0.5 wt-%, usually less than about 0.3 wt-%,calculated as Na₂O.

Fourth, the Y zeolite from the third step is treated further so as toyield Y zeolite having a unit cell size from 24.38 to 24.50 angstrom orpreferably from 24.40 to 24.44 angstrom. In one embodiment, the Yzeolite resulting from the fourth step has an overall silica to aluminamole ratio from 5.0 to 11.0. The treatment of the fourth step cancomprise any of the well known techniques for dealuminating zeolites ingeneral and ultrastable Y zeolite in particular so as to yield thedesired unit cell size and overall silica to alumina mole ratio. Thefourth treatment step may change the unit cell size and/or the frameworksilica to alumina mole ratio, with or without changing the overallsilica to alumina mole ratio. Generally, zeolite dealumination isaccomplished by chemical methods such as treatments with acids, e.g.,HCl, with volatile halides, e.g., SiCl₄, or with chelating agents suchas ethylenediaminetetraacetic acid (EDTA). Another common technique is ahydrothermal treatment of the zeolite in either pure steam or inair/steam mixtures, preferably such as calcining in the presence ofsufficient water vapor (for example, in an atmosphere consistingessentially of steam, and most preferably consisting of steam) so as toyield the desired unit cell size and overall silica to alumina moleratio.

The above-discussed preparation procedure for Y zeolites used in theprocess and composition disclosed herein differs from the procedure forthe Y zeolites taught in U.S. Pat. No. 3,929,672 by the addition of thefourth treatment step. U.S. Pat. No. 3,929,672, which is herebyincorporated herein by reference in its entirety, discloses a method fordealuminating an ultrastable Y zeolite. U.S. Pat. No. 3,929,672 teachesa preparation procedure wherein a sodium Y zeolite is partiallyexchanged with ammonium ions, followed by steam calcination undercontrolled temperature and steam partial pressure, followed by yetanother ammonia exchange and then by an optional calcination step in adry atmosphere. The exchange and steam calcination steps can be repeatedto achieve the desired degree of dealumination and unit cell sizereduction. The zeolites of U.S. Pat. No. 3,929,672 are known under thedesignation Y-84 or LZY-84 commercially available from UOP LLC, DesPlaines, Ill., U.S.A. Y-84 or LZY-84 zeolites may be produced by thefirst three steps just mentioned, but optionally one may include afurther calcination step in a dry atmosphere, e.g., a calcination inwater- and steam-free air, at 482° C. (900° F.) or higher.

The above-discussed preparation procedure for Y zeolites used in theprocess and composition disclosed herein also differs from the procedurefor the Y zeolites taught in U.S. Pat. No. 5,350,501 by differences inthe fourth treatment step. U.S. Pat. No. 5,350,501, which is herebyincorporated herein by reference in its entirety, discloses a fourthstep that involves calcining the resulting zeolite from the thirdtreatment step in the presence of sufficient water vapor (in anatmosphere consisting essentially of steam or consisting of steam) so asto yield a unit cell size below 24.40, and most preferably no more than24.35 angstrom, and with a relatively low sorptive capacity for watervapor. The Y zeolite produced by the four-step procedure in U.S. Pat.No. 5,350,501 is a UHP-Y zeolite, an ultrahydrophobic Y zeolite asdefined in U.S. Pat. No. 5,350,501. The most preferred UHP-Y zeolite inU.S. Pat. No. 5,350,501 is LZ-10 zeolite.

Another group of Y zeolites which may be used in the process andcomposition disclosed herein may be prepared by dealuminating a Yzeolite having an overall silica to alumina mole ratio below about 5 andare described in detail in U. S. Pat. Nos. 4,503,023, 4,597,956, and4,735,928, which are hereby incorporated herein by reference in theirentireties. U.S. Pat. No. 4,503,023 discloses another procedure fordealuminating a Y zeolite involving contacting the Y zeolite with anaqueous solution of a fluorosilicate salt using controlled proportions,temperatures, and pH conditions which avoid aluminum extraction withoutsilicon substitution. U.S. Pat. No. 4,503,023 sets out that thefluorosilicate salt is used as the aluminum extractant and also as thesource of extraneous silicon which is inserted into the Y zeolitestructure in place of the extracted aluminum. The salts have the generalformula:(A)_(2/b)SiF₆wherein A is a metallic or nonmetallic cation other than H⁺ having thevalence “b.” Cations represented by “A” are alkylammonium, NH₄ ⁺, Mg⁺⁺,Li⁺, Na⁺, K⁺, Ba⁺⁺, Cd⁺⁺, Cu⁺⁺, H⁺, Ca⁺⁺, Cs⁺, Fe⁺⁺, Co⁺⁺, Pb⁺⁺, Mn⁺⁺,Rb⁺, Ag⁺, Sr⁺⁺, Ti⁺, and Zn⁺⁺.

A preferred member of this group is known as LZ-210, a zeoliticaluminosilicate molecular sieve available from UOP LLC, Des Plaines,Ill., U.S.A. LZ-210 zeolites and the other zeolites of this group areconveniently prepared from a Y zeolite starting material. In oneembodiment, the LZ-210 zeolite has an overall silica to alumina moleratio from 5.0 to 11.0. The unit cell size is 24.38 to 24.50 angstrom,preferably 24.40 to 24.44 angstrom. The LZ-210 class of zeolites used inthe process and composition disclosed herein have a compositionexpressed in terms of mole ratios of oxides as in the following formula:(0.85-1.1)M_(2/n)O:Al₂O₃:xSiO₂wherein “M” is a cation having the valence “n” and “x” has a value from5.0 to 11.0.

In general, LZ-210 zeolites may be prepared by dealuminating Y-typezeolites using an aqueous solution of a fluorosilicate salt, preferablya solution of ammonium hexafluorosilicate. The dealumination can beaccomplished by placing a Y zeolite, normally but not necessarily anammonium exchanged Y zeolite, into an aqueous reaction medium such as anaqueous solution of ammonium acetate, and slowly adding an aqueoussolution of ammonium fluorosilicate. After the reaction is allowed toproceed, a zeolite having an increased overall silica to alumina moleratio is produced. The magnitude of the increase is dependent at leastin part on the amount of fluorosilicate solution contacted with thezeolite and on the reaction time allowed. Normally, a reaction time ofbetween about 10 and about 24 hours is sufficient for equilibrium to beachieved. The resulting solid product, which can be separated from theaqueous reaction medium by conventional filtration techniques, is a formof LZ-210 zeolite. In some cases this product may be subjected to asteam calcination by methods well known in the art. For instance, theproduct may be contacted with water vapor at a partial pressure of atleast 1.4 kpa(a) (0.2 psi(a)) for a period of between about ¼ to about 3hours at a temperature between 482° C. (900° F.) and about 816° C.(1500° F.) in order to provide greater crystalline stability. In somecases the product of the steam calcination may be subjected to anammonium-exchange by methods well known in the art. For instance, theproduct may be slurried with water after which an ammonium salt is addedto the slurry. The resulting mixture is typically heated for a period ofhours, filtered, and washed with water. Methods of steaming andammonium-exchanging LZ-210 zeolite are described in U.S. Pat. Nos.4,503,023, 4,735,928, and 5,275,720.

The Y zeolites prepared by the above-discussed preparation proceduresand used in the process and composition disclosed herein have theessential X-ray powder diffraction pattern of zeolite Y, and a unit cellsize or dimension a_(o) of 24.38 to 24.50 angstrom, preferably 24.40 to24.44 angstrom. In one embodiment, these Y zeolites have an overallsilica to alumina mole ratio of from 5.0 to 11.0. These Y zeolites mayhave a surface area (BET) of at least about 500 m²/g, at most about 700m²/g, and typically from about 500 to about 650 m²/g. As used herein,surface area means a 20-point surface area determined by standard testmethod UOP874-88, Pore Size Distribution of Porous Substances byNitrogen Adsorption Using a Quantachrome Analyzer, which method isavailable from ASTM International, 100 Barr Harbor Drive, P.O. Box C700,West Conshohocken, Pa., U.S.A.

Another method of increasing the stability and/or acidity of the Yzeolites is by exchanging the Y zeolite with polyvalent metal cations,such as rare earth-containing cations, magnesium cations or calciumcations, or a combination of ammonium ions and polyvalent metal cations,thereby lowering the sodium content until it is as low as the valuesdescribed above after the first or second ammonium exchange steps.Methods of carrying out the ion exchange are well known in the art.

The catalyst used in the process disclosed herein is intended primarilyfor use as a replacement catalyst in existing commercial hydrocrackingunits. Its size and shape is, therefore, preferably similar to those ofconventional commercial catalysts. It is preferably manufactured in theform of a cylindrical extrudate having a diameter of from about 0.8 -3.2mm ( 1/32-⅛ in). The catalyst can however be made in any other desiredform such as a sphere or pellet. The extrudate may be in forms otherthan a cylinder such as the form of a well-known trilobal or other shapewhich has advantages in terms or reduced diffusional distance orpressure drop.

Commercial hydrocracking catalysts contain a number of non-zeoliticmaterials. This is for several reasons such as particle strength, cost,porosity, and performance. The other catalyst components, therefore,make positive contributions to the overall catalyst even if not asactive cracking components. These other components are referred toherein as the support. Some traditional components of the support suchas silica-alumina normally make some contribution to the crackingcapability of the catalyst. In embodiments of the process andcomposition disclosed herein, the catalyst contains a positive amount ofless than about 12, preferably less than about 10 wt-% beta zeolitebased on the combined weight of the beta zeolite, the Y zeolite, and thesupport, all on a dried basis. As used herein the weight on a driedbasis is considered to be the weight after heating in dry air at 500° C.(932° F.) for 6 hours. The catalyst contains at least about 20 wt-%,preferably at least about 30 wt-%, more preferably at least about 40wt-%, and still more preferably from about 45 wt-% to about 60 wt-%, ofY zeolite, based on the combined weight of the beta zeolite, the Yzeolite, and the support, all on a dried basis. Based on the combinedweight of the beta zeolite, the Y zeolite, and the support, all on adried basis, the Y zeolite and beta zeolite content of the catalyst usedin the process disclosed herein is at least about 40 wt-%, preferably atleast about 50 wt-%, even more preferably from about 60 and 80 wt-%,with at least about 50 wt-%, preferably at least about 75 wt-%, evenmore preferably at least about 90 wt-%, and most preferably 100 wt-% ofthe balance being the support.

The remainder of the catalyst particle besides the zeolitic material maybe taken up primarily by conventional hydrocracking materials such asalumina and/or silica-alumina. The presence of silica-alumina helpsachieve the desired performance characteristics of the catalyst. In oneembodiment the catalyst contains at least about 25 wt-% alumina and atleast about 25 wt-% silica-alumina, both based on the combined weight ofthe zeolites and the support, all on a dried basis. In anotherembodiment, the silica-alumina content of the catalyst is above about 40wt-% and the alumina content of the catalyst is above about 35 wt-%,both based on the combined weight of the zeolites and the support, allon a dried basis. However, the alumina is believed to function only as abinder and to not be an active cracking component. The catalyst supportmay contain over about 50 wt-% silica-alumina or over about 50 wt-%alumina based on the weight of the support on a dried basis.Approximately equal amounts of silica-alumina and alumina are used in anembodiment. Other inorganic refractory materials which may be used as asupport in addition to silica-alumina and alumina include for examplesilica, zirconia, titania, boria, and zirconia-alumina. Theseaforementioned support materials may be used alone or in anycombination.

Besides the beta zeolite, the Y zeolite, and other support materials,the subject catalyst contains a metallic hydrogenation component. Thehydrogenation component is preferably provided as one or more basemetals uniformly distributed in the catalyst particle. The hydrogenationcomponent is one or more element components from Groups 6, 9, and 10 ofthe periodic table. Noble metals such as platinum and palladium could beapplied but best results have been obtained with a combination of twobase metals. Specifically, either nickel or cobalt is paired withtungsten or molybdenum, respectively. The preferred composition of themetal hydrogenation component is both nickel and molybdenum or bothnickel and tungsten. The amount of nickel or cobalt is preferablybetween about 2 and about 8 wt-% of the finished catalyst. The amount oftungsten or molybdenum is preferably between about 8 and about 22 wt-%of the finished catalyst. The total amount of a base metal hydrogenationcomponent is from about 10 to about 30 wt-% of the finished catalyst.

The catalyst of the subject process can be formulated using industrystandard techniques. This can, with great generalization, be summarizedas admixing the beta zeolite and the Y zeolite with the other inorganicoxide components and a liquid such as water or a mild acid to form anextrudable dough followed by extrusion through a multihole die plate.The extrudate is collected and preferably calcined at high temperatureto harden the extrudate. The extruded particles are then screened forsize and the hydrogenation components are added as by dip impregnationor the well known incipient wetness technique. If the catalyst containstwo metals in the hydrogenation component these may be addedsequentially or simultaneously. The catalyst particles may be calcinedbetween metal addition steps and again after the metals are added.

In another embodiment, it may be convenient or preferred to combine theporous inorganic refractory oxide, the beta zeolite the Y zeolite, andcompound(s) containing the metal(s), then to comull the combinedmaterials, subsequently to extrude the comulled material, and finally tocalcine the extruded material. In a preferred embodiment, the comullingis effected with ammonium heptamolybdate as the source of molybdenum andnickel nitrate as the source of nickel, with both compounds generallybeing introduced into the combined materials in the form of an aqueoussolution. Other metals can be similarly introduced in dissolved aqueousform or as a salt. Likewise, non-metallic elements, e.g., phosphorus,may be introduced by incorporating a soluble component such asphosphoric acids into the aqueous solution when used.

Yet other methods of preparation are described in U.S. Pat. Nos.5,279,726 and 5,350,501, which are hereby incorporated herein byreference in their entireties.

Catalysts prepared by the above-discussed procedures contain thehydrogenation metals in the oxide form. The oxide form is generallyconverted to the sulfide form for hydrocracking. This can beaccomplished by any of the well known techniques for sulfiding,including ex situ presulfiding prior to loading the catalyst in thehydrocracking reactor, presulfiding after loading the catalyst in thehydrocracking reactor and prior to use at an elevated temperature, andin situ sulfiding, i.e., by using the catalyst in the oxide form tohydrocrack a hydrocarbon feedstock containing sulfur compounds underhydrocracking conditions, including elevated temperature and pressureand the presence of hydrogen.

The hydrocracking process disclosed herein will be operated within thegeneral range of conditions now employed commercially in hydrocrackingprocesses. The operating conditions in many instances are refinery orprocessing unit specific. That is, they are dictated in large part bythe construction and limitations of the existing hydrocracking unit,which normally cannot be changed without significant expense, thecomposition of the feed and the desired products. The inlet temperatureof the catalyst bed should be from about 232° C. (450° F.) to about 454°C. (850° F.), and the inlet pressure should be from about 5171 kpa(g)(750 psi(g)) to about 24132 kpa(g) (3500 psi(g)), and typically fromabout 6895 kpa(g) (1000 psi(g)) to about 24132 kpa(g) (3500 psi(g)). Thefeed stream is admixed with sufficient hydrogen to provide a volumetrichydrogen circulation rate per unit volume of feed of about 168 to 1684normal ltr/ltr measured at 0C. (32° F.) and 101.3 kpa(a) (14.7 psi(a))(1000 to 10000 standard ft^(3/)barrel (SCFB) measured at 15.6° C. (60°F.) and 101.3 kpa(a) (14.7 psi(a))) and passed into one or more reactorscontaining fixed beds of the catalyst. The hydrogen will be primarilyderived from a recycle gas stream which may pass through purificationfacilities for the removal of acid gases although this is not necessary.The hydrogen rich gas admixed with the feed and in one embodiment anyrecycle hydrocarbons will contain at least 90 mol percent hydrogen. Forhydrocracking to produce naphtha the feed rate in terms of LHSV willnormally be within the broad range of about 0.3 to 3.0 hr⁻¹. As usedherein, LHSV means liquid hourly space velocity, which is defined as thevolumetric flow rate of liquid per hour divided by the catalyst volume,where the liquid volume and the catalyst volume are in the samevolumetric units.

The typical feed to the process disclosed herein is a mixture of manydifferent hydrocarbons and coboiling compounds recovered by fractionaldistillation from a crude petroleum. It will normally have componentsthat boil higher than the upper end of the range of the C₆ to 216° C.(420° F.) boiling range. Often it will have a boiling point rangestarting above about 340° C. (644° F.) and ending in one embodimentbelow about 482° C. (900° F.), in another embodiment below about 540° C.(1004° F.), and in a third embodiment below about 565° C. (1049° F.).Such a petroleum derived feed may be a blend of streams produced in arefinery such as atmospheric gas oil, coker gas oil, straight run gasoil, deasphalted gas oil, vacuum gas oil, and FCC cycle oil. A typicalgas oil comprises components that boil in the range of from about 166°C. (330° F.) to about 566° C. (1050° F.). Alternatively, the feed to theprocess disclosed herein can be a single fraction such as a heavy vacuumgas oil. A typical heavy gas oil fraction has a substantial proportionof the hydrocarbon components, usually at least about 80 percent byweight, boiling from about 371° C. (700° F.) to about 566° C. (1050°F.). Synthetic hydrocarbon mixtures such as recovered from shale oil orcoal can also be processed in the subject process. The feed may besubjected to hydrotreating or treated as by solvent extraction prior tobeing passed into the subject process to remove gross amounts of sulfur,nitrogen or other contaminants such as asphaltenes.

The subject process is expected to convert a large portion of the feedto more volatile hydrocarbons such as naphtha boiling rangehydrocarbons. Typical conversion rates vary from about 50 to about 100volume-percent (hereinafter vol-%) depending greatly on the feedcomposition. The conversion rate is between from about 60 to about 90vol-% in an embodiment of the process disclosed herein, from about 70 toabout 90 vol-% in another embodiment, from about 80 and to about 90vol-% in yet another embodiment, and from about 65 to about 75 vol-% instill another embodiment. The effluent of the process will actuallycontain a broad variety of hydrocarbons ranging from methane toessentially unchanged feed hydrocarbons boiling above the boiling rangeof any desired product. The effluent of the process typically passesfrom a reactor containing a catalyst and is usually separated by methodsknown to a person of ordinary skill in the art, including phaseseparation or distillation, to produce a product having any desiredfinal boiling point. The hydrocarbons boiling above the final boilingpoint of any desired product are referred to as unconverted productseven if their boiling point has been reduced to some extent in theprocess. Most unconverted hydrocarbons are recycled to the reaction zonewith a small percentage, e.g. 5 wt-% being removed as a drag stream. Inone embodiment at least 30 wt-%, and preferably at least 50 wt-%, of theeffluent boils below 216° C. (420° F.).

The process and composition disclosed herein can be employed in what arereferred to in the art as single stage and two stage process flows, withor without prior hydrotreating. These terms are used as defined andillustrated in the book titled Hydrocracking Science and Technology, byJ. Scherzer and A. J. Gruia, ISBN 0-8247-9760-4, Marcel Dekker Inc., NewYork, 1996. In a two-stage process the subject catalyst can be employedin either the first or second stage. The catalyst may be preceded by ahydrotreating catalyst in a separate reactor or may be loaded into thesame reactor as a hydrotreating catalyst or a different hydrocrackingcatalyst. An upstream hydrotreating catalyst can be employed as feedpretreatment step or to hydrotreat recycled unconverted materials. Thehydrotreating catalyst can be employed for the specific purpose ofhydrotreating polynuclear aromatic (PNA) compounds to promote theirconversion in subsequent hydrocracking catalyst bed(s). The subjectcatalyst can also be employed in combination with a second, differentcatalyst, such as a catalyst based upon Y zeolite or having primarilyamorphous cracking components.

In some embodiments of the process disclosed herein, the catalyst isemployed with a feed or in a configuration that the feed passing throughthe catalyst is a raw feed or resembles a raw feed. The sulfur contentof crude oil, and hence the feed to this process, varies greatlydepending on its source. As used herein, a raw feed is intended to referto a feed which has not been hydrotreated or which still containsorganic sulfur compounds which result in a sulfur level above 1000wt-ppm or which still contains organic nitrogen compounds that result ina nitrogen level above 100 wt-ppm (0.01 wt-%).

In other embodiments of the process disclosed herein, the catalyst isused with a feed that has been hydrotreated. Persons of ordinary skillin the art of hydrocarbon processing know and can practice hydrotreatingof a raw feed to produce a hydrotreated feed to be charged to theprocess disclosed herein. Although the sulfur level of the feed may bebetween 500 and 1000 wt-ppm, the sulfur level of the hydrotreated feedis less than 500 wt-ppm in one embodiment of the process disclosedherein and from 5 to 500 wt-ppm in another embodiment. The nitrogenlevel of the hydrotreated feed is less than 100 wt-ppm in one embodimentand from 1 to 100 wt-ppm in another embodiment.

All references herein to the groups of elements of the periodic tableare to the IUPAC “New Notation” on the Periodic Table of the Elements inthe inside front cover of the book titled CRC Handbook of Chemistry andPhysics, ISBN 0-8493-0480-6, CRC Press, Boca Raton, Fla., U.S.A.,80^(th) Edition, 1999-2000. All references herein to boiling points areto boiling points as determined by ASTM D2887, Standard Test Method forBoiling Range Distribution of Petroleum Fractions by Gas Chromatography,which method is available from ASTM International.

The following examples are provided for illustrative purposes and not tolimit the process and composition as defined in the claims.

EXAMPLE I Catalyst 1

A sample of commercially-available steamed and ammonium-exchanged LZ-210was obtained from UOP LLC. The LZ-210 had an overall silica to alumina(SiO₂ to Al₂O₃) mole ratio of from 8.0 to 10.0 and a unit cell size of24.42 angstrom. Catalyst 1 was prepared by comulling a mixture of theLZ-210, beta zeolite having an overall silica to alumina (SiO₂ to Al₂O₃)mole ratio of 23.8 and an SF₆ adsorption capacity of 29 wt-%,HNO₃-peptized Catapal™ C boehmite alumina (Catapal C alumina isavailable from Sasol North America, Inc., Houston, Tx., U.S.A.),sufficient nickel nitrate to provide 5 wt-% nickel (calculated as NiO)in the final catalyst and sufficient ammonium heptamolybdate to provide15 wt-% molybdenum (calculated as MoO₃) in the final catalyst. Thecomulled mixture was extruded into 1.6 mm ( 1/16 in) diametercylindrical particles of between 3.2 mm (⅛ in) and 12.7 mm (½ in) inlength, dried, and then calcined at about 500° C. (932° F.) for aminimum of 90 minutes. The resulting catalyst contained nickel andmolybdenum in the proportions above specified. The catalyst contained 45wt-% LZ-210 zeolite, 5 wt-% beta zeolite, and 50 wt-% alumina based onthe combined weight of the zeolites and the support all on a driedbasis.

Catalyst 2

Catalyst 2 was prepared similarly to Catalyst 1 except more of theLZ-210 and more of the beta zeolite was used. The resulting catalystcontained nickel and molybdenum in the proportions above specified forCatalyst 1. The catalyst contained 60 wt-% LZ-210 zeolite, 10 wt-% betazeolite, and 40 wt-% alumina based on the combined weight of thezeolites and the support all on a dried basis.

Catalyst 3 (comparative)

Catalyst 3 was prepared similarly to Catalyst 1 except the LZ-210 wasused in place of the beta zeolite. The resulting catalyst contained thenickel and molybdenum in the proportions above specified for Catalyst 1.The catalyst contained 50 wt-% LZ-210 zeolite and 50 wt-% alumina basedon the combined weight of the zeolites and the support all on a driedbasis.

Catalyst 4 (comparative)

Catalyst 4 was prepared similarly to Catalyst 2 except the LZ-210 wasused in place of the beta zeolite. The resulting catalyst containednickel and molybdenum in the proportions above specified for Catalyst 2.The catalyst contained 70 wt-% LZ-210 zeolite and 30 wt-% alumina basedon the combined weight of the zeolites and the support all on a driedbasis.

Catalyst 5 (comparative)

Catalyst 5 was prepared similarly to Catalyst 1 except more of theLZ-210 and more of the beta zeolite was used. The resulting catalystcontained nickel and molybdenum in the proportions above specified forCatalyst 1. The catalyst contained 50 wt-% LZ-210 zeolite, 12.5 wt-%beta zeolite, and 37.5 wt-% alumina based on the combined weight of thezeolites and the support all on a dried basis.

EXAMPLE II

Each of the above-described five catalysts was pre-sulfided by passing agas stream consisting of 10 vol-% H₂S and the balance H₂ through a bedof the catalyst at a temperature initially of about 149° C. (300° F.)and slowly raised to 413° C. (775° F.) and held at the temperature forabout 6 hours.

The five catalysts were compared for hydrocracking activity andselectivity (i.e., product yields) in simulated first stage testing.Specifically, the five catalysts were separately tested forhydrocracking a hydrotreated coker gas oil feed having a specificgravity of 0.8647 at 15.6° C. (60° F.) (API gravity of 32.15°), aninitial boiling point of 94° C. (202° F.), a 5 wt-% boiling point of170° C. (338° F.), a final boiling point of 440° C. (824° F.), and a 50wt-% boiling point of 297° C. (566° F.), with about 15 wt-% boilingbelow 216° C. (420° F.).

Each catalyst was tested for simulated first stage operation by passingthe feedstock through a laboratory size reactor at a LHSV of 1.5 hr⁻¹, atotal pressure of 9997 kpa(g) (1450 psi(g)), and a volumetric hydrogenfeed rate per unit volume of feed of 1684 normal ltr/ltr measured at 0°C. (32° F.) and 101.3 kPa(a) (14.7 psi(a)) (10000 SCFB measured at 15.6°C. (60° F.) and 101.3 kpa(a) (14.7 psi(a))). Sufficient di-tert-butyldisulfide was added to the feed to provide 1.70 wt-% sulfur and therebyto simulate a hydrogen sulfide-containing atmosphere as it exists incommercial first stage hydrocracking reactors. In addition, sufficientcyclohexylamine was added to the feed to provide 2640 wt-ppm nitrogenand thereby to simulate an ammonia-containing atmosphere as it exists incommercial first stage hydrocracking reactors. The temperatureconditions were adjusted as necessary to maintain about a 65 wt-% netconversion to materials boiling below 216° C. (420° F.), over the courseof 100 hours. Net conversion is the effluent boiling below 216° C. (420°F.) as a percentage of the feed minus the percentage of the feed boilingbelow 216° C. (420° F.). At the end of the 100 hours, the temperaturerequired to maintain the 65 wt-% net conversion was recorded, and theactivity and selectivity of each catalyst relative to a commercialreference were calculated. These data are summarized in the Table. Theactivity and yield data for each catalyst are entered as the differencebetween the actual value for activity or yield of the catalyst minus theactual value for activity or yield obtained with the reference. The morenegative the value for activity, the more active is the catalyst.

TABLE Composition, wt-% of zeolite and support on a Catalyst designationdried basis 1 2 3 4 5 Zeolites 50 70 50 70 62.5 Y zeolite 45 60 50 70 50Beta zeolite 5 10 — — 12.5 Support 50 30 50 30 37.5 Y:Beta ratio, 9 6 NANA 4 weight/weight Relative activity, −4 −7 −2 −3 −6 ° C. (° F.) (−8)(−12) (−3) (−6) (−11) Relative C₆-216° C. +0.7 +0.6 +0.8 +0.7 −0.1(C₆-420° F.) fraction yield, wt-% NA = Not Applicable

As shown in the Table, the catalytic activity of Catalyst 1, a catalystdisclosed herein containing a combination of beta zeolite and Y zeolitehaving a unit cell size of 24.45 angstrom, is substantially greater thanthat of the reference in both activity and yield. Catalyst 1 is 4.4° C.(8° F.) more active than the reference, and also shows a significantadvantage over the reference with respect to yield of the C₆ to 216° C.(420° F.) boiling fraction. The catalytic activity of Catalyst 1 is alsogreater than that of Catalyst 3, which is a catalyst that has the sametotal amount of zeolite as Catalyst 1 but has no beta zeolite, at nearlythe same yield of the C₆ to 216° C. (420° F.) boiling fraction.

The Table also shows that the catalytic activity of Catalyst 2, anothercatalyst disclosed herein containing a combination of beta zeolite and Yzeolite having a unit cell size of 24.45 angstrom, is substantiallygreater than that of the reference in both activity and yield. Catalyst2 is 6.7° C. (12° F.) more active than the reference, and also shows asignificant advantage over the reference with respect to yield of the C₆to 216° C. (420° F.) boiling fraction. The catalytic activity ofCatalyst 2 is also greater than that of Catalyst 4, which is a catalystthat has the same total amount of zeolite as Catalyst 2 but has no betazeolite, at nearly the same yield of the C₆ to 216° C. (420° F.) boilingfraction.

Catalyst 5 shows that an excess of beta zeolite, when used incombination with a Y zeolite having a unit cell size of 24.45 angstrom,gives nearly the same yield of the C₆ to 216° C. (420° F.) boilingfraction as the reference.

1. A process for hydrocracking a hydrocarbon feedstock which comprisescontacting the feedstock at a temperature from about 232° C. to about454° C. and at a pressure from about 5171 kPa(g) to about 24132 kPa(g)in the presence of hydrogen with a catalyst comprising a hydrogenationcomponent, a beta zeolite, and a Y zeolite having a unit cell size from24.38 to 24.50 angstrom, the catalyst having a weight ratio of Y zeoliteto beta zeolite of from 5 to 12 on a dried basis.
 2. The process ofclaim 1 wherein the weight ratio of the Y zeolite to the beta zeolite isfrom 6 to 9 on a dried basis.
 3. The process of claim 1 wherein the unitcell size is from 24.40 to 24.44 angstrom.
 4. The process of claim 1wherein the catalyst comprises a support and contains at least about 20wt-% Y zeolite based on the combined weight of the beta zeolite, the Yzeolite, and the support all on a dried basis.
 5. The process of claim 1wherein the catalyst comprises a support and contains a positive amountof less than about 12 wt-% beta zeolite based on the combined weight ofthe beta zeolite, the Y zeolite, and the support all on a dried basis.6. The process of claim 1 wherein the catalyst comprises a support andcontains at least about 40 wt-% of beta zeolite and Y zeolite based onthe combined weight of the beta zeolite, the Y zeolite, and the supportall on a dried basis.
 7. The process of claim 1 wherein the Y zeolitehas an overall silica to alumina mole ratio from 5.0 to 11.0.
 8. Theprocess of claim 1 wherein the Y zeolite is prepared by a processcomprising the steps of: a) partially ammonium exchanging a sodium Yzeolite; b) calcining the zeolite resulting from step (a) in thepresence of water vapor; c) contacting the zeolite resulting from step(b) with a fluorosilicate salt in the form of an aqueous solution; andd) calcining the zeolite resulting from step (c) in the presence ofwater vapor.
 9. The process of claim 8 wherein the zeolite resultingfrom step (d) has a BET surface area of from about 500 to about 700m²/g.
 10. The process of claim 1 wherein the Y zeolite is an LZ-210zeolite.
 11. The process of claim 1 wherein the Y zeolite is prepared bya process comprising the steps of: a) partially ammonium exchanging asodium Y zeolite; b) calcining the zeolite resulting from step (a) inthe presence of water vapor; c) ammonium exchanging the zeoliteresulting from step (b); and d) calcining the zeolite resulting fromstep (c) in the presence of water vapor.
 12. The process of claim 1wherein the beta zeolite has an overall silica to alumina mole ratio ofless than 30.0.
 13. The process of claim 1 wherein the beta zeolite hasa SF₆ adsorption capacity of at least 25 wt-%.
 14. The process of claim1 wherein the hydrogenation component is selected from the groupconsisting of an IUPAC Group 6 component, an IUPAC Group 9 component,and a IUPAC Group 10 component.
 15. The process of claim 14 wherein thehydrogenation component is selected from the group consisting ofmolybdenum, tungsten, nickel, cobalt, and the oxides and sulfidesthereof.
 16. A hydrocracking process comprising contacting a hydrocarbonfeedstock with a catalyst at a temperature between about 232° C. andabout 454° C. and at a pressure between about 5171 kPa(g) and about24132 kPa(g) in the presence of hydrogen so as to produce an effluent oflower average boiling point than the hydrocarbon feedstock, the catalystcomprising one or more hydrogenation components in combination with asupport comprising an inorganic refractory oxide, zeolite beta in a formcatalytically active for cracking hydrocarbons, and a Y zeolitecatalytically active for cracking hydrocarbons, the Y zeolite having aunit cell size of from 24.38 to 24.50 angstrom, the catalyst having aweight ratio of zeolite beta to Y zeolite of from 5 to
 12. 17. Thehydrocracking process of claim 16 wherein the Y zeolite has an overallsilica to alumina mole ratio of from 5.0 to 11.0.
 18. The hydrocrackingprocess of claim 16 wherein at least 30 wt-% of the effluent boils below216° C.
 19. The hydrocracking process of claim 16 wherein at least 50wt-% of the effluent boils below 216° C.